Direct coal liquefaction process

ABSTRACT

A direct coal liquefaction process capable of producing unexpectedly high levels of C5/650° F. product, which process employs a relatively high ratio of solvent plus bottoms product recycle to feed coal.

CROSS-REFERENCE TO RELATED APPLICATIONS

This is a continuation-in-part of U.S. Ser. No. 14/147,542 filed Jan. 5,2014.

FIELD OF THE INVENTION

The present invention relates to a direct coal liquefaction processcapable of producing unexpectedly high levels of C5/650° F. product,which process employs a relatively high ratio of solvent plus bottomsproduct recycle to feed coal.

BACKGROUND OF THE INVENTION

From time to time, increases in cost and future shortages of petroleumoften leads to increased interest in coal as a fuel source given thevast, easily accessible deposits of coal that exist in several parts ofthe world. Various processes have been proposed for converting coal toliquid and gaseous fuel products, including transportation fuels, and insome cases, to other products such as lubricants and chemicals. Problemsthat have hampered the commercial liquefaction of coal include therelatively low thermal efficiency of indirect coal-to-liquids (ICTL)conversion methods, such as Fischer Tropsch (FT) synthesis andmethanol-to-liquids (MTL) conversion, as well as high water use and CO₂emissions.

Direct coal liquefaction (DCL) methods have typically involved heatingthe coal in the presence of a hydrogen donor solvent, and optionally acatalyst, in a hydrogen containing atmosphere to a temperature in therange of about 700° to 950° F. This results in a break-down of the coalstructure into free radicals that are quenched to produce liquidproducts. The catalyst typically contained finely divided iron,molybdenum, or mixtures thereof. One source of the molybdenum catalystis via in situ formation from a phosphomolybdic acid (PMA) precursor.

Conventional direct coal liquefaction process units typically requirepassing spent solvent to a hydrotreater for preparing a hydrogen donorsolvent that is fed back to the input of the DCL unit to act as asolvent and to increase conversion in the liquefaction process.Unfortunately, solvent hydrotreating requires separation of at least afaction of the solvent, additional equipment, and additionalhydrogen-rich treat gas. Addition of an external solvent hydrotreaterincreases the required investment and decreases thermal efficiency. Thisis an important reason why low donor solvent-to-coal ratios (typically1.2 to 1.5) are used in coal liquefaction process units using a solventhydrotreater.

In addition, hydrotreating reduces the viscosity and lowers thearomatics content of the solvent, which reduces its ability to suspendash in the slurry and its compatibility with coal. The reduction inability to suspend ash results in an increased likelihood of solidsbuildup, deposits, or plugging of high pressure feed pumps, transferlines, heat exchangers, furnace tubes, and reactors. Hydrotreatedsolvent also has a higher cracking rate during liquefaction. This canresult in the liquefaction process producing insufficient productsolvent for recycle for preparing a slurry with the feed coal. For sucha situation, an external source of solvent (such as coal tar) isrequired.

Use of donor solvents also results in higher gas hold-up in theliquefaction reactor, which in turn, requires a larger reactor volume toachieve adequate coal residence time in the reactor. Further, highrecycle gas rates are also required because treat gas must be providedfor both the solvent hydrotreater and the liquefaction reactor.

In order to reduce gas hold-up and avoid solids build up, ebullated bedreactors are preferred. Because of the high liquid recycle in ebullatedbed reactors, the reactors are substantially fully back-mixed, whichresults in an increase in reactor volume versus plug flow reactors.

As much as 50% of the heat of reaction is moved from the liquefactionreactor to the solvent hydrotreater. Thus, additional fuel must be firedfor preheating of the feed to liquefaction and for the solventhydrotreater. This results in lower thermal efficiency.

A further issue limiting the application of DCL processes is that lowerquality coals, such as those having inertinite content higher than about12 vol. % have been considered unsuitable for use as a DCL feedstock.High inertinite coals are found in many parts of the world, includingthe United States and China. Many of these coals, such as that in theOrdos basin in China have inertinite contents of more than 25% and havea low ratio of atomic hydrogen to carbon (H/C). Such coals havehistorically been unacceptably more difficult to liquefy by DCL thanhigher quality coals that have a substantially low vitrinite content.

Okada reported in a paper entitled: “Possible Impacts of Coal Propertieson the Coal Conversion Technology”, Coal Science, J. A. Pajares and J.M. D. Tascon, 1995 Elsevier Science, that oil yield from autoclaveexperiments were inversely proportional to the inertinite content ofcoal for coals of similar rank. Oil yield ranged from about 67 wt % fora zero inertinite content coal to about 40 wt % for a coal containingabout 60 vol % inertinite. He concluded that high inertinite coals, suchas that found in the Ordos Basin, are not suitable for direct coalliquefaction.

Also, in the paper “Study on Coal Liquefaction Characteristics ofChinese Coals”, Fuel 81 (2002) 1551-1557, Wasaka published results on 53runs on 27 coals that were made in a 0.1 t/d pilot plant test program.The program specifically focused on direct liquefaction of highinertinite content coals. At constant liquefaction operating conditions,conversion decreased with increasing inertinite content coals.

Further U.S. Pat. No. 7,763,167 to Zhang et al discloses a DCL processthat utilized an iron-containing catalyst and an externally hydrotreateddonor solvent. The donor solvent was produced in a suspended bed using aforced circulation reactor (ebullated bed). Although they obtained anoil yield, they did not disclose the boiling point of the oil product.

While the art contains various conventional DCL processes to convertvarious coals to liquids, there remains a need in the art for a DCLprocess that is able to achieve unexpectedly high yields of C5/650° F.boiling range products, even with relatively high inertinite-containingcoals.

SUMMARY OF THE INVENTION

In accordance with the present invention there is provided a process forthe direct liquefaction of coal, which process is conducted in theabsence of added carbon monoxide, and is performed in a coalliquefaction process plant comprising: a slurry mixing zone, apreheating zone, a liquefaction zone, a separation zone capable ofseparating a gaseous product stream from a liquid/solids product stream,an atmospheric fractionation zone, and a vacuum fractionation zone,which process comprises:

a) introducing into said slurry mixing zone:

-   -   i) coal having an average particle size of about 75 microns to        about 600 microns and a moisture content from about 1 to about 4        wt. %;    -   ii) non-donor solvent from said vacuum fractionation zone and        non-donor bottoms from said atmospheric fractionation zone,        wherein the ratio of non-donor solvent plus non-donor bottoms to        coal is from about 2.5 to 1 to about 4 to 1:    -   iii) a molybdenum-containing catalyst provided at a make-up rate        that is equivalent to about 50 wppm to about 2 wt. % molybdenum        on a moisture and ash free (MAF) feed coal basis, wherein the        resulting slurry is at a temperature from about 200° F. to about        600° F.;

b) conducting said slurry at a pressure from about 1500 psig to about3000 psig and an effective amount of treat gas containing at least about80 vol. % hydrogen, to said preheating zone wherein it is heated to atemperature of about 650° F.;

c) conducting said heated slurry to said liquefaction zone wherein it isreacted at a temperature from about 700° F. to about 950° F. therebyproducing reaction products;

d) conducting said reaction products to a separation zone wherein agaseous product stream is separated from a liquid/solids product stream;

e) conducting said liquid/solids stream to said atmosphericfractionation zone, wherein it is fractionated to result in a C1 to C4gaseous fraction, a C5/650° F. fraction, and a non-donor 650 F+ bottomsfraction;

f) conducting an effective portion of non-donor 650° F.+ bottomsfraction from said atmospheric fractionation zone to said vacuumfractionation zone thereby resulting in a 1000° F.+ fraction and anon-donor 650° F. to 1000° F. solvent fraction from the vacuumfractionator; and leaving a remaining portion of said non-donor 650° F.+fraction from said atmospheric fractionator;

g) recycling the remaining portion of said non-donor 650° F.+ fractionfrom said atmospheric fractionation zone to the slurry mixing zone; and

h) recycling at least a portion of said non-donor 650° F. to 1000° F.solvent fraction from said vacuum fractionation zone to said slurrymixing zone.

In a preferred embodiment of the present invention the ratio ofnon-donor solvent plus non-donor bottoms to coal is about 3:1 to 4:1.

In another preferred embodiment of the present invention the ratio ofnon-donor solvent plus non-donor bottoms to coal is about 3:1 to about3.5:1.

In yet another preferred embodiment of the present invention the amountof inertinite in said coal is from about 7 to 14 vol. %.

In still another preferred embodiment of the present invention theamount of inertinite in said coal is greater than about 20 vol. %.

BRIEF DESCRIPTION OF THE FIGURES

FIG. 1 hereof is a flow diagram of a preferred process scheme for thepractice of the present invention.

FIG. 2 hereof is a bar chart showing the impact of total coal conversionversus the ratio of recycle non-donor solvent plus recycle non-donorbottoms to coal for the examples herein.

FIG. 3 hereof is a bar chart of yield of C5/650° F. product versus theratio of recycle non-donor solvent plus recycle non-donor bottoms tocoal for the examples herein.

DETAILED DESCRIPTION OF THE INVENTION

It has unexpectedly been found by the inventors hereof that a dramaticand unexpected increase in C5/650° F. liquid product, along with highthermal efficiency, are achieved by the practice of the presentinvention. This dramatic and unexpected increase can also be achievedwith any bituminous or sub-bituminous coal, even using a coal having ahigh inertinite content. For example, it is well known in the art thatthe Wyoming coals (Rawhide and Wyodak) typically have an inertinitecontent from about 8 to 11 vol. %. Coals from other regions of thePowder River Basin in Wyoming can have inertinite contents of up to 20vol. %. There are also Chinese coals having very high inertinitecontent. For example, Yulin coal from the province of Shaanxi has ainertinite content of about 30 vol. % and Guojiawan coal has aninertinite content of about 46 vol. %. It is well known in the art thathigh inertinite content coals are difficult to convert by direct coalliquefaction processes.

One key to the unexpected results obtained by the direct coalliquefaction process of the present invention is the use of a high ratioof non-donor recycle solvent and non-donor recycle bottoms to coal. Thatis, high ratios of: i) the 650° F. to 1000° F. fraction from the vacuumfractionator zone, sometimes referred to herein as “solvent”, and ii)the 650° F.+ fraction from the atmospheric fractionation zone, sometimesreferred to herein as “bottoms”. Typically, the upper range of the 650°F.+ fraction from the atmospheric fractionator will be about 700° F. Theratios are in excess of those conventionally practiced commercial sizepilot plants and process plants by direct coal liquefaction processes.Such plants can typically process at least about 75 lbs of coal per day.This dramatic increase in C5/650° F. product yield of the presentinvention is particularly present in the instant micro-catalytic directcoal liquefaction process. That is wherein a finely dispersed molybdenumor iron catalysts is used. It is preferred that the catalyst of thepresent invention be comprised of finely dispersed molybdenum, and thata non-donor product recycle stream produced in liquefaction process beused. The ratio of such recycle solvent and bottoms product to coal atthe input to the liquefaction zone, on a moisture free weight basis, isat least about 2.5:1, preferably about 3:1, also about 3.5:1, as well asabout 4:1.

The term “non-donor” as used herein means that the recycle solvent andstreams have not been processed in a hydrotreater to partiallyhydrogenate multi-ring aromatic compounds in the stream to producecompounds that can donate hydrogen during liquefaction. Surprisingly,increasing the ratio of the recycled non-donor stream to coal in theinstant process does not increase the flow rate of the recycled streamand fresh coal to liquefaction for a given rate of product generation.Instead, surprisingly, less coal is required. Although the recyclestream increases relative to coal, the total feed to liquefactionremains substantially the same. The net impact of higher recycle andlower coal rate is a reduction of energy required in the slurrypreheating zone and the size of the vacuum fractionator. Hence,investment and energy requirements are reduced for the liquefactionsection of the instant liquefaction process.

Referring now to FIG. 1 hereof, there is presented a preferredembodiment of the present DCL process. A coal feed is dried and crushed,preferably to an average particle size of from about 30 to 200 mesh,corresponding from about 75 microns to about 600 microns. The particlesize reduction can be performed in any suitable mill for reducingparticle to the sized set forth about, but it is preferred that a gasswept roller mill 201 be used. It is also preferred that the moisturecontent of the milled coal be from about 1 to 4%. The resulting crushedand dried coal is fed to mixing tank 203 where it is mixed with recyclestreams derived from both the atmospheric and vacuum fractionators. Acatalyst precursor is introduced, preferably is in the form of anaqueous solution of phosphomolybdic acid (PMA) in an amount that isequivalent to adding between 50 wppm and 2 wt. % molybdenum on amoisture and ash free (MAF) coal feed basis. Slurry mix tank 203, whichis preferably operated at a temperature from about 200° F. to about 600°F., preferably 300° F. to 600° F., and more preferably from about 300°F. to about 500° F. From the slurry mix tank, the catalyst containingslurry is delivered to slurry pump 205. The selection of the appropriatemixing and temperature conditions is based on experimental workquantifying the rheological properties of the specific slurry blendbeing processed, and is easily within the skill of those having at leastordinary skill in the art.

Most of the remaining moisture in the coal is driven off in the mixingtank due to hot atmospheric fractionator bottoms feeding into saidmixing tank. Residual moisture, as well as any entrained volatiles, arecondensed out as sour water (not shown). The coal in the slurry leavingmixing tank 203 has about a 0.1 to 1.0% moisture content. The slurryformed by the coal and the recycle streams is pumped from mixing tank203 wherein the pressure is raised to about 1,500 to 3,000 psig,preferably from about 2000 to 3000 psig, by slurry pumping system 205.The resulting high-pressure slurry may be preheated in a heat exchanger(not shown), mixed with a treat gas comprised of recycled and makeuptreat gas containing at least about 80 vol. % hydrogen, then furtherheated to a temperature of about 600° F. to about 700° F. in slurrypreheating furnace 207. The coal slurry and hydrogen mixture is fed tothe input of the first reactor of the series connected liquefactionreactors 209, 211 and 213 at between about 600° F. to 700° F. and 1,500to 3,000 psig. Reactors 209, 211 and 213 are preferably up-flow tubularvessels. The total length of the three reactors is from about 40 to 200feet. The temperature rises from one reactor stage to the next as aresult of the highly exothermic coal liquefaction reactions. In order tomaintain the maximum temperature in each stage below about 700 to 950°F., preferably from about 800° F. to about 900° F. It is preferred thata portion of the hydrogen-containing treat gas is preferably injectedbetween reactor stages. The hydrogen partial pressure in each stage ispreferably maintained at a minimum of about 1,000 to 2,000 psig.

The effluent from the last reactor is separated into a gaseous streamand a liquid/solid stream. The liquid/solid stream is let down inpressure separation and cooling zone 215, also sometimes referred toherein as the separation zone. The gaseous stream is cooled to condenseout liquid vapors, such as H₂O, naphtha, distillate, and solvent. Theremaining gas is then processed to remove H₂S and CO₂.

Most of the processed gas is then sent to a hydrogen recovery system,not shown, for further processing by conventional means to recover atleast a fraction of the hydrogen contained therein, which is thenrecycled to be mixed with the coal slurry. Any remaining processed gascan be purged to prevent buildup of light ends in the recycle loop.Hydrogen recovered therefrom can be used in any downstreamhydro-processing upgrading system.

The resulting depressurized liquid/solid stream and the hydrocarbonscondensed in separation and cooling stage 215 are passed to atmosphericfractionation zone 219 where they are separated into light ends andliquid fractions. The liquid fraction is separated into a light endsfraction, a C5/650° F. minus liquid fraction, and a 650° F.+ fraction.The light ends are processed to recover hydrogen and C1-C4 hydrocarbonsthat can be used for fuel gas and other purposes.

An effective amount or portion of the of the 650° F.+ fraction fromatmospheric fractionation zone 219 is passed to the vacuum fractionationzone 221 wherein it is separated into a non-donor 650° F. to 1000° F.solvent fraction and a 1000° F.+ fraction. By effective amount orportion we mean that amount required to purge ash from the system. Theremaining portion of 650° F.+ stream from the atmospheric fractionationzone and the 650° F. to 1000° F. solvent fraction from the vacuumfractionation zone are recycled to slurry mix tank, also sometimesreferred to herein as the slurry mixing zone 203.

At least a fraction of the 1000° F.+ fraction from vacuum fractionator221 is sent to be gasified by partial oxidation zone 223 to generatehydrogen for use in the present liquefaction process. A portion of thecoal from gas sweep mill 201 is preferably fed to partial oxidation zone223 to produce additional hydrogen. Alternatively, instead of thepartial oxidation zone 223, the 1000° F.+ bottoms fraction from vacuumfractionator 221 can be processed in a Circulating Fluid Bed boiler, acement plant, or sold as a feed for asphalt paving or for electrodemanufacture.

Hydrogen for liquefaction and upgrading can also be produced by anysuitable technology. For example, Steam Methane Reforming of a streamsuch as natural gas, shale gas, or coal mine methane, which is wellknown in the art, can be used. Such a technology is utilized worldwidein refineries and offered by many commercial vendors such asHaldor-Topsoe. Also, catalysts useful in DCL processes of the presentinvention include those disclosed in U.S. Pat. Nos. 4,077,867, 4,196,072and 4,561,964, the disclosures of which are hereby incorporated hereinby reference in their entirety.

A preferred embodiment of the liquefaction process of the presentinvention combines several elements that contribute to maximum premiumC5/650° F. fuels product production and maximum thermal efficiency.These include: the recycle of a non-donor streams, preferably includingatmospheric fractionator bottoms, to maintain a ratio of the recyclestream to coal at the input to reactors 209, 211, 213 that is at least2.5:1 on a moisture free weight basis, preferably between 3.0:1 and3.5:1, as well as 4:1; and the use of a micro-catalyst in the form offinely dispersed molybdenum. Also, use of 650° F.+ recycle, and multipleslurry reactors in series contribute to the benefits of the instantprocess. It will be understood that the expression (S+B) which issometimes used herein is meant to mean (Solvent+Bottoms), both of whichare non-donor recycle streams.

Use of a micro-catalyst, which is either a compound of molybdenum oriron, more preferably molybdenum, and added at 100 to 1,000 wppm, morepreferably 100 to 500 wppm, and most preferably 100 to 300 wppm,eliminates several disadvantages associated with the use of a donorsolvent such as required by prior DCL systems. Firstly, energy is lostduring the preparation of the donor solvent. Secondly, energy isrequired to preheat the donor solvent in the solvent hydrotreater andhydrogen must be compressed and circulated around the hydrotreater.Thirdly, the heat release during partial hydrogenation of the donorsolvent is lost during cooling prior to separation of hydrogen forrecycle. In comparison, all of the heat release occurs in the presentprocess in the liquefaction reactors during operation with a 650° F.+recycle stream, which minimizes the preheat requirement prior toliquefaction. These factors contribute to the higher thermal efficiencyof the microcatalytic coal liquefaction process. Moreover, the use of amicrocatalyst and the consequent elimination of the need for a donorsolvent also eliminates the need for an expensive solvent hydrotreaterto generate the donor solvent, thereby substantially reducing thecapital cost of the system. It also permits the use of coals havingsubstantially higher ash contents, from 6 to 20 wt % or more on amoisture free basis, and the recycle of a substantially higher portionof 650° F.+ than were possible with donor solvent systems. Examples ofmicrocatalysts and their method of preparation are described in U.S.Pat. No. 4,226,742, the contents of which are incorporated herein byreference in their entirety.

Further, the 650° F.+ fraction from atmospheric fractionator 219 and the650° F. to 1000° F.− fraction from vacuum fractionator 221, as thenon-donor streams being recycled to the slurry mixing tank 203 providespreheat for the coal and solvent in slurry mix tank 203. This preferablyraises the temperature in the mixing 203 tank to about 200° F. to 600°F., preferably 300° F. to 600° F., more preferably about 300° F. to 500°F., and most preferably about 400 to 500° F. This further reduces theenergy requirement for preheating the slurry prior to liquefaction. Asignificant portion of the micro-catalyst is entrained in the 600 to700° F.+ fraction recycled from the atmospheric fractionator 219, sothat recycling a larger portion of such fraction increases the catalystconcentration in the liquefaction reactors, thereby decreasing therequirement for the addition of fresh catalyst precursor and increasingthe conversion efficiency of the process.

Use of the non-donor 600° F. to 700° F.+ stream, preferably 630° F. to670° F.+, and more preferably a 650° F.+, process derived recyclesolvent in the DCL process of the present invention reduces cracking,relative to use of a donor solvent, and thus produces a 650° F.− producthaving a greater fraction of diesel and less light gases and naphtha.The 650° F.− product can be selectively upgraded by conventionalprocessing to finished products in fixed bed upgrading reactors.

The use of two to four, more preferably three slurry reactors in seriesapproaches a plug flow reactor and hence has as little as two thirds ofthe required volume of one or two ebullated bed reactors such as used insome prior DCL systems. Since all of the heat is released in the threeliquefaction reactors, the temperature profile can be maintained tomaximize selectivity to liquids. Thus, the use of three seriallyconnected reactors are preferred. Operation of the initial reactor at asomewhat lower temperature has been reported in previous patents as aroute to increase conversion and liquid yields. An exemplary process forupgrading the liquid product of the DCL reactors is disclosed in U.S.Pat. No. 5,198,099, the disclosure of which is hereby incorporatedherein by reference in its entirety.

The diesel product, after upgrading, will have a Cetane number ofbetween approximately 42 and 47, depending upon cut points of theproduct and aromatics content. Specific gravity of the product will alsovary between 0.83 and 0.90. A higher Cetane Number is required for Euro4 diesel, thus a Fischer-Tropsch facility producing a 70-75 CetaneNumber diesel blend stock may be added to the plant operating inaccordance with the present invention. The gasoline produced byupgrading the relevant portion of product of the process of the presentinvention will meet all current gasoline specifications, or can beupgraded to a Research Octane of 106 if desired. This will permit theblending of the low octane naphtha into the gasoline pool whilemaintaining adequate octane for the blended fuel. Also, the upgradingprocess can also be operated to maximize the production of jet fuel orgasoline. The jet fuel produced will meet all Military JP-8specifications.

Continuous, integrated pilot plants are utilized to define commercialcoal conversion and product yields. These facilities typically includedistillation towers for recycle of solvent, bottoms, and catalystcontained in the bottoms. It generally takes seven or eight days forcatalyst concentration, coal conversion, composition of streams, andyields to equilibrate and provide real world data for commercial coalliquefaction plants. Small batch reactors, such as tubing bombs,mini-bombs, and autoclaves, that are typically used for examples in mostcoal liquefaction patent applications, merely provide basic informationto researchers. They are not capable of providing reliable data for usein designing a commercial direct coal liquefaction plant. Further, suchsmall reactors are not designed to be continuous reactors, but representa batch experiment with catalyst, stream composition, and relative ratesthat are not equilibrated because of their designed short run times. Inaddition, because of their small size, extraction with solvents such ascyclohexane or THF (tetrahydrofuran) are utilized for determiningconversion rather than distillation (fractionation) which is a veryimportant step in a commercial plant to identify a range of products.Typically, only total conversion is capable of being determined with useof such small once-through reactors. Thus, the present invention appliesonly to those direct coal liquefaction plants containing both anatmospheric fractionator and a vacuum fractionator. It is preferred thatthe size of the liquefaction plant be at least about 75 lb per day coalfeed, and more preferably at least about 240 lb per day of coal. It willbe understood that the terms atmospheric distillation, atmosphericfractionator, atmospheric distillation tower, and atmospheric pipe stillcan be used interchangeably herein. Also, the terms vacuum distillation,vacuum fractionator, vacuum distillation tower, and vacuum pipe stillcan be used interchangeably herein as well.

The following examples provide coal conversion and yield data for eithera 75 lb/day or 240 lb/day pilot plant. Both units have been designed ormodified to operate as a continuous, integrated pilot plant thatutilizes both atmospheric and vacuum distillation for separation ofproducts and recycle streams.

Data presented in the following examples was obtained at liquefactiontemperatures of 780° F. to 880° F. and operating pressures of 2,000 to2,500 psig. Catalysts used include (1) iron, (2) iron and molybdenum, or(3) molybdenum alone. Summaries of the data obtained in the followingexamples are presented in Table 1 and in FIGS. 2 and 3 hereof

Example 1 (Comparative)

Wasaka et al, discloses in “Study on Coal Liquefaction Characteristicsof Chinese Coals; Fuel 81 (2002); 1551-1557, the direct coalliquefaction of 27 Chinese coals having inertinite contents between 0.8vol. % and 46.1 vol. %. Solvent was recycled at a ratio solvent to coalof 1.5 for each run with each coal. All runs were performed in a 240lb/day pilot plant in China.

The pilot plant was comprised of a slurry preparation and feed section,a liquefaction section, and a product separation and withdrawal section.A slurry of catalyst, donor solvent and coal was prepared in the slurrypreparation and feed section and was blended with hydrogen from ahydrogen feed section, then sent to a liquefaction section. It wasretained in the liquefaction section for about 1 hour at a temperatureof about 878° F. (470° C.) and a pressure of about 2466 pia.

As an example, Guojiawan was one of the coals tested by Wasaka et al.This coal contained 46.1 vol % inertinite. The donor solvent washydrotreated to produce a hydrogen donor solvent prior to preparation ofthe slurry with the feed coal. A catalyst make-up rate of up to 3 wt. %Fe₂O₃ was used as a catalyst in addition to the donor solvent. Both theiron catalyst and the donor solvent are known routes for increasing coalconversion, but both are unacceptable with the present process becausethe solvent used in the present invention is a non-donor solvent and thecatalyst is molybdenum.

The liquefaction products were analyzed and it was found that totalconversion was about 70 wt. % and it was estimated that the C5/650 Freaction product would be about 37.5 wt. %, which are both shown inTable 1 and FIGS. 2 and 3 hereof.

Examples 2 and 3 (Comparative)

Two coal liquefaction experiments were performed on a high inertinitecontent (about 29.9 wt. %), Yulin coal from Inner Mongolia. The recyclesolvent+bottoms to coal ratio used was 1.2 for Example 2 and 1.5 forExample 3.

Both examples were performed in a 240 lb/day continuous coalliquefaction pilot plant in Beijing, China. The pilot plant was modifiedto have substantially the same configuration as the 75 lb/day continuouscoal liquefaction pilot plant hereof for examples 4 and 5 hereof. Forexample, the 75 lb/day pilot plant contained a slurry preparation andfeed section, a liquefaction section, and a product separation andwithdrawal section which were also included in the 240 lb/day plant. Forthis example, three liquefaction reactors were operated in series attemperatures of 788° F., 842° F., and 842° F. respectively. Pressure foreach reactor was 2,500 psig with 120 minutes nominal run time, and acatalyst make-up of 300 wppm molybdenum for Example 4 and 1000 wppmmolybdenum for Example 5.

The liquefaction products were analyzed and it was found that total coalconversion was only 57.2 wt. % for Example 2, and 57.8 wt. % for Example3. C5/650° F. yields was 41 wt % and 41.2 wt. % for Examples 2 and 3,respectively.

The low total coal conversion resulting from these two examples isconsistent with the higher inertinite content of the Yulin coal. It isconventional wisdom in the art that it is difficult to convert asubstantial fraction of high inertinite coal in DCL (Direct CoalLiquefaction) processes.

Example 4 and 5 (Comparative)

Two coal liquefaction experiments were performed at a recycle(solvent+bottoms)/coal ratio of 1.65.

Both experiments were performed with Rawhide coal, which had ainertinite content of about 9.5 vol. %, in the same 75 lb/day continuouscoal liquefaction pilot plant that included fractionation of theresulting liquid product and recycle of solvent and bottoms. Reactorconditions were: 801° F. and 842° F. in two stages of liquefaction; eachat a pressure of 2500 psig; 50 minutes nominal run time, and a make-uprate of mixed catalyst containing 100 wppm molybdenum and about 1 wt %Fe₂O₃.

The resulting liquefaction products were analyzed and it was found thattotal coal conversion was 89.3 wt. % for Example 4 and 88.6 wt % forExample 5. C5/650° F. yield was 46.1 for Example 4 and 44.2 wt. % forExample 5.

Examples 6 to 11 (Comparative)

Six experiments were performed using the same Rawhide coal used inExamples 4 and 5 hereof but at recycle (solvent+ bottoms)/C of 1.7,1.71, 1.74, 1.74, 1.83 and 1.91, respectively.

These experiments were also performed in the same 75 lb/day continuouscoal liquefaction pilot plant used for Examples 4 and 5 hereof.Liquefaction temperatures ranged from about 807° F. to about 836° F. inthe first stage and 842° F. in the second stage. For each example, thepressure was 2500 psig. Nominal residence time ranged from 41 to 49minutes. The liquefaction catalysts make-up rate was 100 wppm moly and0.25 to 1.03 wt. % FeO₃.

The resulting liquefaction products were analyzed and it was found thattotal coal conversion for the six Examples was found to be 85.6, 87.0,86.9, 85.2, 85.7, and 88.6 wt % respectively. The C5/650° F. yield forthe six Examples was found to be 44.0, 43.5, 41.9, 42.4, 44.4, and 48.3wt. % respectively.

Example 12 to 14 (Comparative)

Three experiments were performed using Wyodak coal having an inertinitecontent of about 10 vol. %. All three examples used a recycle (solvent+bottoms)/Coal ratio of 2.0.

These experiments were performed in the 75 lb/day pilot plant asprevious mentioned in prior examples hereof. The liquefactionsconditions for each experiment were: average liquefaction temperature of797° F.; pressure of 1969 psig; a nominal residence time of about 144minutes; and a molybdenum make-up rate 100 wppm.

The resulting liquefaction products were analyzed and it was found thatthe total coal conversion for example 12 to 14 was 77.5, 81.4, and 75.9wt. % respectively. The C5/650° F. yield was found to be 47.1, 44.2, and45.1 wt. %, respectively.

Examples 15 and 16 (Comparative)

Two experiments were performed using the Wyodak coal of Examples 12 to14 hereof at a (solvent+ bottoms)/Coal ratio of 2.0 but at higherliquefaction temperatures.

These experiments were performed in the 75 lb/day pilot plant discussedin prior examples hereof. The average liquefaction temperature for eachexample was 841° F.; at a pressure of 1994 psig for example 15 and 2412psig for Example 16; a nominal residence time of 48 minutes for Example15 and 70 minutes for Example 16. Both examples used a molybdenummake-up rate of 100 wppm on MAF coal.

The resulting liquefaction products were analyzed and it was found thatthe total coal conversion for Example 15 was 78.2 wt. % and 76.1 wt. %for Example 16. C5/650° F. yield was found to be 47.8 wt. % to 43.5 wt.% for Examples 15 and 16 respectively.

Example 17 (Example of this Invention)

A coal liquefaction experiment was performed using the same Yulin coalused in Examples 2 and 3 hereof but a recycle (solvent+ bottoms)/Coalratio of 3.

The same 240 lb/day pilot plant and conditions were utilized as inExamples 2 and 3 hereof except for the higher (Solvent+Bottoms)/Coalratio of at least 3. The molybdenum catalyst make-up rate was 300 wppm.

The resulting liquefaction products were analyzed and it was found thattotal coal conversion was unexpectedly found to be 86.5 wt. % versus 57wt. % as in Examples 2 and 3 hereof. It was also unexpectedly found thatC5/650° F. yield increased to 58.3 wt. %. If 650/700° F. is included,the total liquid yield would be 66.4 wt. %.

All coal conversion and C5/650° F. yield data from the 17 Examples issummarized in Table 1 and in FIGS. 1 and 2 hereof. As indicated in theTable and Figures:

Coal conversion for the low inertinite Rawhide and Wyodak coals wasbetween the high 70's and high 80's depending upon operating conditionsand catalyst concentration and were independent of recycle (solvent+bottoms)/Coal ratio between zero and 2.0. For the high inertinite Yulincoal, total coal conversion was below 60 wt. % at a(Solvent+Bottoms)/Coal ratio of 1.2 to 1.5. This low coal conversion hasbeen reported by other researchers for high inertinite content coals.Increasing the (Solvent+Bottoms)/Coal ratio to 3.0 for Yulin coalunexpectedly increased conversion to substantially the same level asthat for low inertinite coals.

In addition to increasing coal conversion with high inertinite contentcoals, operating at high recycle (solvent+ bottoms)/Coal ratio of 3:1also unexpectedly increased the direct coal liquefaction yield ofC5/650° F. in the product. This lower boiling product can be readilyconverted to gasoline, diesel, and jet fuel using fixed bed upgradingreactors which are common to today's refineries. This evidences that itis possible to even convert a high inertinite content coal to anunexpectedly high yield of C5/650° F. product.

TABLE 1 Pilot Plant Volume 1000 Exam- Feedrate % wt % F.- ple lb/dayCoal Inertinite (S + B)/C C5/650 Conv 1 240 Guojiawan 46.1 1.5 37.5 est70 2 240 Yulin 29.9 1.2 41.03 57.23 3 240 Yulin 29.9 1.5 41.15 57.75 475 Rawhide 9.5 1.65 46.09 89.26 5 75 Rawhide 9.5 1.65 44.21 88.6 6 75Rawhide 9.5 1.7 44.04 85.62 7 75 Rawhide 9.5 1.71 43.46 87.01 8 75Rawhide 9.5 1.74 41.91 86.88 9 75 Rawhide 9.5 1.74 42.35 85.19 10 75Rawhide 9.5 1.83 44.41 85.74 11 75 Rawhide 9.5 1.92 48.31 88.56 12 75Wyodak 10 2.00 47.1 77.5 13 75 Wyodak 10 2.00 44.2 81.4 14 75 Wyodak 102.00 45.1 75.9 15 75 Wyodak 10 2.00 47.8 78.2 16 75 Wyodak 10 2.00 43.576.1 17 240 Yulin 29.9 3.00 58.27 86.54

What is claimed is:
 1. A process for the direct liquefaction of coal,which process is conducted in the absence of added carbon monoxide, andis performed in a coal liquefaction process plant comprising: a slurrymixing zone, a preheating zone, a liquefaction zone, a separation zonecapable of separating a gaseous product stream from a liquid/solidsproduct stream, an atmospheric fractionation zone, and a vacuumfractionation zone, which process comprises: a) introducing into saidslurry mixing zone: i) coal having an average particle size of about 75microns to about 600 microns and a moisture content from about 1 toabout 4 wt. %; ii) effective amount of recycle non-donor solvent fromsaid vacuum fractionation zone and non-donor bottoms from saidatmospheric fractionation zone, wherein the ratio of non-donor solventplus non-donor bottoms to coal is from about 2.5 to 1 to about 4 to 1:iii) a molybdenum-containing microcatalyst provided at a make-up ratethat is equivalent to about 50 wppm to about 2 wt. % molybdenum on amoisture and ash free (MAF) feed coal basis; b) conducting the resultingslurry from step a) above at a pressure from about 1,500 psig to about3000 psig and an effective amount of a treat gas containing at least 80vol. % hydrogen, to said preheating zone wherein it is heated from atemperature from about 200° F. to 600° F. to a temperature of about 650°F.; c) conducting said heated slurry to said liquefaction zone whereinit is reacted at a temperature from about 700° F. to about 950° F.thereby producing reaction products; d) conducting said reactionproducts to a separation zone wherein a gaseous product stream isseparated from a liquid/solids product stream; e) conducting saidliquid/solids stream to said atmospheric fractionation zone, wherein itis fractionated to result in a C1 to C4 gaseous fraction, a C5/650° F.fraction, and a 650 F+ bottoms fraction; f) conducting an effectiveportion of non-donor 650° F.+ bottoms fraction from said atmosphericfractionation zone to said vacuum fractionation zone thereby resultingin a 1000° F.+ fraction and a non-donor 650° F. to 1000° F. solventfraction from the vacuum fractionator; and leaving a remaining portionof said non-donor 650° F.+ fraction from said atmospheric fractionator;g) recycling the remaining portion of said non-donor 650° F.+ fractionfrom said atmospheric fractionation zone to the slurry mixing zone; andh) recycling at least a portion of said non-donor 650° F. to 1000° F.solvent fraction from said vacuum fractionation zone to said slurrymixing zone.
 2. The process of claim 1 wherein the ratio of non-donorsolvent plus non-donor bottoms to coal is about 3:1 to 4:1.
 3. Theprocess of claim 2 wherein the ratio of non-donor solvent plus non-donorbottoms to coal is about 3:1 to about 3.5:1.
 4. The process of claim 1wherein the amount of inertinite in said coal is from about 7 to 14 vol.%.
 5. The process of claim 1 wherein the amount of inertinite in saidcoal is greater than about 20 vol. %.
 6. The process of claim 1 whereinsaid treat gas contains at least about 70 vol. % hydrogen.
 7. Theprocess of claim 6 wherein said treat gas contains at least about 80vol. % hydrogen.
 8. The process of claim 1 wherein the pressure is fromabout 2000 psig to about 3000 psig.
 9. The process of claim 1 whereinthe liquefaction temperature if from about 800° F. to about 900° F. 10.The process of claim 7 wherein the pressure is from about 2000 psig toabout 3000 psig and the liquefaction temperature is from about 800° F.to about 900° F.